Process of hydrofining high nitrogen hydrocarbons followed by catalytic cracking with zeolitic aluminosilicates



April 14, 1970 ANNESSER ET AL 3,506,568

PROCESS OF HYDROFINING HIGH NITROGEN HYDROCARBONS FOLLOWED B Y CATALYTIC CRACKING WITH ZEOLITIC ALUMINOSILICATES Original Filq @pril 26, 1966 2 Sheets-Sheet 1 AMORPHOUS CATALYST ZEOLITIC T ALUMINOSILICATE CATALYST NITROGEN, PPM.

O 5 0 v 5 O 6 5 5 4 4 FIG.

INVENTORS RICHARD J. ANNESSER HE/NZ R WEBER WI'ORNEYS April. 14, 1.970

I rnocsss OF mnnomnme men NITROGEN nmnocmaonsromownn BY CATALYTIC CRACKING WITH ZEQLITIC ALUMINOSILICATES Original Filed April as. 1966 Sv lA-NNESSER E L 3,506,568

2 Sheets-Sheet 2 ZEOLITIC ALUMINOSILICATE CATALYST AMORPHOUS CATALYST NITROGEN, PPM.

FIG. 2

INVENTORS RICHARD J. ANNESSER HE/NZ WEBER United States Patent 3,506,568 PROCESS OF HYDROFINING HIGH NITROGEN HYDROCARBONS FOLLOWED BY CATA- LYTIC CRACKING WITH ZEOLITIC ALUMI- NOSILICATES Richard J. Annesser and Heinz P. Weber, Berkeley, Calif.,

assignors to Chevron Research Company, San Francisco, Califi, a corporation of Delaware Continuation of application Ser. No. 553,115, Apr. 26, 1966. This application Jan. 10, 1969, Ser. No. 790,357 Int. Cl. Cg 23/02 US. Cl. 208-89 3 Claims ABSTRACT OF THE DISCLOSURE Hydrofining a high boiling hydrocarbon feed containing substantial amounts of organic nitrogen to reduce the organic nitrogen content prior to catalytic cracking of the feed with a crystalline zeolitic aluminosilicate in a catalytic cracking zone.

CROSS-REFERENCE This application is a continuation of application Ser. No. 553,115, filed May 26, 1966 now abandoned.

The present invention relates to catalytic cracking with catalysts comprising crystalline zeolitic aluminosilicates. More particularly, the present invention is concerned with hydrofining hydrocarbon feeds containing high levels of nitrogen impurities followed by catalytic cracking of the hydrofined feeds in the presence of crystalline zeolitic aluminosilicate catalysts.

A principal process in the petroleum industry by which high boiling hydrocarbons are converted to lower boiling products, including gasoline, is catalytic cracking. Catalytic cracking is to be distinguished from hydrocracking. Hydrocracking involves hydrogenation at temperatures high enough for cracking to occur, whereas catalytic cracking primarily involves cracking in the absence of hydrogen, thereby preventing significant hydrogenation. A distinctive feature of catalytic cracking is the high octane quality of the gasoline produced, resulting from the presence of high concentrations of branched chain parafiin hydrocarbons and olefin hydrocarbons. Catalytic cracking also yields highly unsaturated C and C fractions, and high concentrations of isobutane.

Catalytic cracking of hydrocarbon feeds has generally been performed using catalysts comprising amorphous silica-alumina composites, e.g., natural clays, or synthetic silica-alumina cogels. It is known that organic nitrogen in the feed is detrimental to such amorphous catalysts and characteristically deactivates the catalysts by, for example, increasing coke production, thereby reducing the yield and quality of the products, and in particular the yield and quality of gasoline. In general, the reduction in yield is proportional to the amount of organic nitrogen in the cracking feed. The petroleum industry has attempted to overcome the detrimental effects of nitrogen by removing the nitrogen compounds through catalytic hydrofining of the feed prior to catalytic cracking. However, hydrofining or hydrodenitrification has only been partially successful. The removal of nitrogen has proven to be very difficult. For example, in order to remove substantial quantities of nitrogen from a hydrocarbon feed, it is often necessary to use high pressures of up to 3000 "ice p.s.i.g. or even higher in some cases. The difiiculty in hydrofining is especially significant when utilizing heavy feeds, such as vacuum gas oils and residues. Moreover, the hydrofining operation is generally very expensive due to the pressures involved, the cost of the equipment and the necessity for an available source of hydrogen. As a result, hydrofining of catalytic cracking feeds to remove nitrogen has not gained commercial acceptance since the benefits obtained have not been commensurate with the expenses involved. To justify hydrofining of catalytic cracking feeds, additional significant advantages would necessarily have to be shown.

The advent of zeolitic aluminosilicates as new catalytic cracking catalysts was considered by some as the solution to the problem of the deactivating effect of organic nitrogen impurities in catalytic cracking feedstocks. A principal advantage of zeolitic aluminosilicates was the high conversion obtainable with nitrogen containing feeds. US. Patent 2,962,435 discloses that the presence of zeolitic aluminosilicates in catalytic cracking catalyst mixtures effectively neutralizes the catalyst mixtures against the adverse effects of nitrogen compounds present in the feed. Hence, the use of zeolitic aluminosilicates for cracking catalysts was supposed to make hydrofining of nitrogen containing catalytic cracking feeds even less economical than when amorphous catalysts were used for catalytic cracking.

It is now found, however, that hydrofining of catalytic cracking feedstocks having nitrogen contents above about 2500 ppm. is economically justifiable when the hydrofining reduces the nitrogen to below about 2500 p.p.m. and when the subsequent catalytic cracking is accomplished in the presence of catalysts comprising zeolitic aluminosilicates. It is indeed surprising to find hydrofining so beneficial in view of the expenses involved in hydrofining and especially in view of the tolerance zeolitic aluminosilicates were supposed to have towards nitrogen. In accordance with the present invention, unexpected advantages in the conversion of a high boiling hydrocarbon feed having a nitrogen content above about 2500 ppm. to lower boiling products, including gasoline, are obtained by a process which comprises contacting said feed with a hydrofining catalyst in the presence of hydrogen under hydrofining conditions. adapted to reduce the nitrogen content below about 2500 p.p.m., then contacting the hydrofined fraction with a catalyst comprising a crystalline zeolitic aluminosilicate in which the pore dimensions are at least 6 A. under catalytic cracking conditions in a catalytic cracking zone.

The present invention will be more fully explained and understood by reference to the figures in the attached drawings which are graphs illustrating the catalytic cracking conversion to products as a function of the nitrogen content (degree of hydrofining of a hydrocarbon gas oil) at a constant coke yield of 5 weight percent.

The graphs in the figures were obtained 'by converting to a constant coke basis the results of numerous catalytic cracking runs involving hydrofined feeds of different nitrogen levels. It is found that although catalytic cracking with zeolitic aluminosilicates results in higher conversion of the feeds to lower boiling products than amorphous catalysts, the higher conversion is accompanied by higher coke yields. Since most commercial catalytic cracking processes are limited by their coke-burning capacity, conversion may be limited thereby. Thus,

3 comparison of catalysts on the basis of constant coke yield, as is done in the figures, is highly desirable. At constant coke yield it is found that with feeds having a high nitrogen content (i.e., greater than about 2500 p.p.m.), zeolitic aluminosilicate catalysts are no better in terms of conversion of the feeds than amorphous catalysts. Surprisingly though, zeolitic aluminosilicates are superior to amorphous catalysts with feeds having low nitrogen contents (i.e., below about 2500 p.p.m.). Thus a distinct advantage is gained in hydrofining feeds to below 2500 p.p.m. nitrogen. This was unexpected because the alleged nitrogen tolerance of aluminosilicates was supposed to make hydrofining less attractive for improving catalytic cracking feeds.

Thus, as a specific embodiment, by use of the invention improved results can be obtained in a catalytic cracking process comprising a reaction zone wherein a hydrocarbon feedstock is converted and a regeneration zone wherein coke is burned from the catalyst, between which zones cracking catalyst is continuously circulated, wherein the feedstock available for catalytic cracking contains above 2500 p.p.m. nitrogen, and wherein limitations in the coke burning capacity of said regeneration zone impose an upper limit on the conversion attainable in said reaction zone, thereby preventing effective utilization in said process of a cracking catalyst comprising a crystalline zeolitic aluminosilicate to obtain increased conversion because coke formation accompanying cracking of the feedstock containing above 2500 p.p.m. nitrogen with an aluminosilicate-containing catalyst is greater than with an aluminosilicate-free catalyst at the same reaction conditions and because conversion obtainable with an aluminosilicate-containing catalyst is not substantially greater than with an aluminosilicatefree catalyst at reaction conditions yielding the same amount of coke. The invention comprises hydrofining said feedstock to reduce its nitrogen content to below 2500 p.p.m., passing the hydrofined feedstock to said cracking process, and using therein a cracking catalyst comprising a crystalline zeolitic aluminosilicate, which makes possible obtaining increased conversion of said feedstock without increased coke formation.

In general, the process of the present invention is preferably applied to hydrocarbon feeds having an initial boiling point above about 450 F. and an end point below about 1100 F. The nitrogen content of the feeds for purposes of the present invention should preferably be greater than 2500 p.p.m. The feeds can contain sulfur and metal contaminants along with nitrogen. Suitable feeds include virgin crudes, vacuum distillation residues, gas oils, and solvent deasphalted oils. These feeds may be derived from petroleum crude oils, shale oils, tar sand oils, coal hydrogenation and carbonization products and the like.

For purposes of the present invention the nitrogen content of the feedstock is reduced in the hydrofining operation from above about 2500 p.p.m. to below about 2500 p.p.m. The advantages of hydrofining prior to catalytic cracking with a catalyst comprising a zeolitic aluminosilicate begin to be realized when the hydrofined fraction contains less than about 2500 p.p.m. nitrogen. The greater the nitrogen removal the greater the benefit derived from hydrofining prior to catalytic cracking of the feed. Hence, it is preferred to reduce the nitrogen content to below about 2000 p.p.m. and more preferably below about 1000 p.p.m.

The hydrofining is conducted under conditions of temperature, pressure, hydrogen flow rate and liquid hourly space velocity in the reactor correlated to provide the desired degree of nitrogen removal. Higher temperatures, pressures, and hydrogen flow rates are used when treating the higher boiling feedstocks and those containing greater amounts of organic nitrogen.

The temperature has a large influence on the rate of conversion of the nitrogen compounds, and is adjusted upwards to maintain the proper degree of hydrofining as the catalyst ages or is deactivated through prolonged use. The temperature should be in the range 500 to 850 F. and preferably in the range 600 to 800 F. At temperatures below 500 F. the rate of hydrofining, or nitrogen removal, is too low for practical purposes, whereas at temperatures above about 850 F. substantial cracking of the feed occurs, and coke formation tends to increase markedly. The temperature used will also depend on the activity of the hydrofining catalyst, higher temperatures being used with a less active catalyst.

The pressure should be maintained within the range 400 to 4000 p.s.i.g. and preferably within the range from 800 to 3000 p.s.i.g. Elevated pressures advantageously influence the rate and extent of hydrofining, as well as extend the catalyst activity and life. However, higher pressures increase the cost of the hydrofining operation.

The liquid hourly space velocity (LHSV), that is, the flow of hydrocarbon feed relative to the catalyst, will generally be in the range 0.2 to 10 and preferably within the range of 0.3 to 5. In general, the nitrogen compounds found in high boiling hydrocarbon feeds are considered more resistant to hydrofining than those found in lower boiling feeds. Hence, the space velocity is generally lower for higher boiling feeds, but depends significantly on the other hydrofining conditions as well as the desired degree of nitrogen removal.

The flow of hydrogen into the reactor is maintained above about 500 s.c.f./bbl. of feed and preferably in the range 1,000 to 10,000 s.c.f./bbl. and more preferably, 1,000 to 4,000 s.c.f./bbl. More generally, at least suflicient hydrogen is provided to supply that consumed in the conversion of nitrogen compounds to ammonia and compensate for incidental hydrogenation of unsaturates and oxygen and sulfur compounds, while maintaining a significant excess of hydrogen partial pressure. Hydrogen can be added to the feed prior to introduction of the feed into the reactor; or the hydrogen can be added separately to the reactor. The hydrogen consumption will generally be within the range to 2,000 s.c.f./bbl. of feed depending on the properties of the hydrocarbon feed and the other hydrofining conditions used. Excess hydrogen is removed from the treated oil, and preferably purified and recycled to the reaction zone.

In the hydrofining zone the feed plus added hydrogen is contacted with any suitable hydrofining catalyst. Suitable hydrofining catalysts generally comprise the Group VIII metals, their oxides and/or sulfides thereof mixed with Group VI-B metals, their oxides and/or sulfides thereof. The metal composites may be used in the undiluted form but preferably exist in combination with a support. Suitable carriers or supports are the inorganic oxides, for example, alumina, silica, zicronia, titania, bauxite, magnesia, fullers earth, and combinations thereof. The metal content on a support preferably ranges between about 2 percent to 25 percent by weight. Suitable hydrofining catalysts contemplated for use in the present invention include cobalt oxide and molybdenum oxide on silica-alumina; sulfided nickel and tungsten on alumina; and nickel-molybdenum on alumina. A particularly good catalyst is nickel and molybdenum on a silica-alumina support.

The form in which the hydrofining catalyst is used will depend on the type of process involved in the hydrofining operation, that is whether the process involves a fixed lbed, moving bed, of fluid operation. Generally, the catalyst will exist in beads, tablets or extruded pellets for use in fixed bed or moving bed operations, and in powder form for use in fluid operations. If the catalyst maintains high activity over protracted periods of use, the hydrofining is preferably carried out using a fixed bed of catalyst in a reactor. Catalyst regeneration can be periodically accomplished by subjecting the catalyst to an oxygen-containing atmosphere at elevated temperatures to remove carbon deposits formed during extended use.

Following the hydrofining operation, the hydrofined feed may be treated so as to remove any contaminants, such as ammonia, which may be present. Removal of ammonia may be accomplished, for example, by injecting water or acidified water into the hydrofined feed and passing the resulting mixture into a separator operating under such conditions that a water phase containing essentially all the ammonia present in the hydrofined feed can be removed. Further purification of the hydrofined feed can be accomplished in a stripper or a distillation column. For purposes of the present invention, however, it is not considered essential to treat hydrofined feed to remove the contaminants produced during hydrofining. Hence the hydrofined, feed can generally be catalytically cracked in the presence of the zeolitic aluminosilicate catalyst without intervening purification.

At least a portion and preferably all the hydrofined feed containing less than about 2500 p.p.m. nitrogen can be cracked in the presence of a catalyst comprising a crystalline zeolitic aluminosilicate. Both the natural and synthetic zeolitic aluminosilicates may be used for purposes of the present invention. Crystalline zeolitic aluminosilicates comprise aluminosilicate cage structures in which alumina and silica tetrahedra are intimately connected with each other in an open three dimensional network. The tetrahedra are cross-linked by the sharing of oxygen atoms. In general, the spaces between the tetrahedra are occupied by water molecules prior to dehydration. De-

hydration results in crystals interlaced With channels or pores of molecular dimensions which channels or pores selectively limit the size and shape of foreign substances that can be adsorbed. Thus, the crystalline zeolitic aluminosilicates are often referred to as molecular sieves. In the hydrated form the aluminosilicates can be represented by the basic formula:

wherein M is a cation which balances the negative electrovalence of the tetrahedra; n represents the valence of the cation; w, the moles of SiO and y, the moles of water. In general, a particular type of crystalline zeolitic aluminosilicate will have values of w and y that fall in a definite range. The cation, M, may be any of a number of ions, such as, for example, the alkali metal ions, the alkaline earth ions, and the rare earth ions. The cations may be mono-, di-, or trivalent. The zeolitic cations may be replaced one with another by suitable exchange techniques. The replacement of the zeolitic cations with other cations, as, for example, the replacement of sodium cations with calcium cations, generally does not induce appreciable changes in the anionic framework.

The aluminosilicates which find use for purposes of the present invention possess relatively well defined pore structures. The exact type of aluminosilicate is relatively unimportant as long as the pore structure comprises openings characterized by pore dimensions greater than '6 A. and, in particular, uniform pore diameters of between approximatelyo A. and A. The uniform pore structure wherein the pores are greater than 6 A. permit bydrocarbons access to the reactive sites of the catalyst. Generally, in order to obtain aluminosilicates of the necessary pore diameters, the silica to alumina ratio in the crystalline form should be greater than about 2. Appropriate zeolitic aluminosilicates which find use in the present invention are the natural faujasites; synthesized zeolite X described in U.S. Patent 2,882,244; and zeolite Y described in U.S. Patent3,130,007. Zeolite Y is generally more stable under catalytic cracking conditions and hence is the preferable form of the aluminosilicates.

Generally, the crystalline zeolitic aluminosilicate catalyst will not contain metal hydrogenating components.

However, a number of other ions may be incorporated into the aluminosilicate structure, as for example the alkali metals, the alkaline earths and the rare earths. It is preferred to maintain the sodium content of the zeolitic aluminosilicate below about 10 weight percent based on the oxide. The hydrogen form of the zeolitic aluminosilicate can also be used.

The zeolitic aluminosilicate can be employed directly as a catalyst or it can be combined with other suitable catalytic materials, as, for example, silica-alumina or silica-rriagnesia. Furthermore, the zeolitic aluminosilicate can be mixed with a support or binder to provide beneficial properties such as increased compactibility and attrition resistance. The particular chemical composition of the sup port or binder is not critical. It is, however, necessary that the support or binder employed be thermally stable under the conditions at which the cracking is carried out. The support or binder may be catalytically inert or possess catalytic activity. Such materials include by way of example kieselguhr, bauxite and various clays. The mixture can be prepared by a variety of methods, as, for example, by physically mixing and then compressing the composite, or by coprecipitation, or cogellation.

Reaction conditions depend on the type of catalytic cracking process employed, whether fixed bed, moving bed, or fluid. Furthermore, the cracking conditions depend on the nature of the feedstock, Whether highly paraffinic or aromatic, etc., and upon the nitrogen content. In general, the reaction conditions, such as temperature, pressure, and liquid hourly space velocity are correlated to provide the yield and nature of products desired.

The temperature in the catalytic cracking operation should lie within the range from 700 to 1200 F. and preferably within the range 800 to 1000 F. generally, increasing the temperature increases the amount of cracking or the conversion of feed to lower boiling products.

The appropriate pressure can be from subatmospheric to several atmospheres. Preferably the pressure will lie within the range 5-100 p.s.i.g. and more preferably, 520 p.s.i.g. The pressure has little eifect on the rate of cracking although it affects the contact time. Moreover, increasing the pressure generally reduces the octane quality of the gasoline product and increases the production of coke at a given conversion.

The liquid hourly space velocity is preferably maintained within the range 0.5-20 and more preferably from 1-5. The catalyst to oil ratio should be maintained between about 1 to 20 on a weight basis and preferably from 2 to 10. The catalyst to oil ratio depends on the type of process used, whether a fluid, moving bed or fixed bed, and generally higher catalyst to oil ratios are used for fluid operations. Increasing the catalyst to oil ratio normally reduces the extent of catalyst deactivation from coke production, and increases the conversion of the feed to lower boiling products.

The present process can be conducted in either fixed bed, moving bed, or fluid catalyst systems. Because of the coke laydown on the catalyst and the necessity of regenerating the catalyst periodically it is preferred to employ a contacting system wherein regeneration can be accomplished without discontinuing the flow of feed to the reaction zone. A particularly preferred contacting system is one involving a fluid catalyst. In this operation a finely divided solid catalyst, for example powder, is continuously recycled between a reaction zone and a separate regeneration zone. In each zone the catalyst is maintained ina fluidized state that behaves much like a liquid in the reactor. The feed is continuously contacted with freshly regenerated catalyst and the hydrocarbon products are removed from the reactor. The coked catalyst is continuously removed from the reactor and passed to a regenerator where it is contacted with an oxygencontaining atmosphere to burn the coke and regenerate the catalyst. The regenerated catalyst is then returned to the reaction zone.

The present process may be more fully understood in terms of the following example.

7 EXAMPLE A California gas oil, boiling between about 500 F. and 985 F., and having the specifications set forth in Table I, was used in the hydrofining-catalytic sequence The untreated gas oil was contacted under hydrofining conditions and in the presence of hydrogen with a catalyst comprising nickel and molybdenum on a silica-alumina support. Several hydrofining runs were made at different reaction conditions to obtain treated or hydrofined gas oils of different levels of nitrogen content. For each run the hydrofined oil removed from the reactor was transmitted to a high pressure separator where hydrogen gas (containing some hydrogen sulfide and light hydrocarbon gases) was removed. The removed gas was recycled to the reactor along with added hydrogen. Conventional techniques were used to purify the hydrofined oil of The untreated gas oil feed and the hydrofined fractions were contacted under catalytic cracking conditions with a catalyst comprising a zeolitic aluminosilicate having pore dimensions greater than approximately 6 A. The catalyst also contained about 17 weight percent silicaalumina thoroughly admixed with the zeolitic aluminosilicate. For purposes of comparison, catalytic cracking of the untreated gas oil and the hydrofined fractions was performed in the presence of an amorphous catalytic cracking catalyst. The amorphous catalyst was composed of the mixed refractory oxides of silica and alumina; the catalyst did not contain any zeolitic aluminosilicates. The silica-alumina catalyst is very effective for the catalytic cracking of gas oils and the like to other lower boiling hydrocarbons, including gasoline, and is used extensively in many present day catalytic cracking processes. The untreated gas oil feed as well as hydrofined portions thereof were catalytically cracked in the presence of the amorphous catalyst under conditions similar to those used with the zeolitic aluminosilicate containing catalyst. The operating conditions as well as the important product yields of representative runs are tabulated in Table III-A for the amorphous catalyst and in Table III-B for the zeolitic aluminosilicate containing catalyst.

The total conversion tabulated in Tables III-A and III-B was determined from the difference between the volume of the feed and the volume of the materials boiling above 430 F.; thus total conversion values include not only the amount of gasoline boiling range materials produced, but also the light gases and coke. The space velocity was measured in terms of weight of feed per hour per weight of catalyst (w./hr./w.).

TABLE III-A [Amorphous catalyst] Run No 1 2 3 4 5 6 7 8 9 Nitrogen in feed, p.p.rn 4, 650 4, 650 4, 650 3, 200 2, 500 2, 500 1, 700 470 470 Operating conditions:

Reactor tempcrature, F 910 950 950 910 910 910 910 910 910 Reactor pressure, .s.1.g 10 10 10 10 10 1O 10 i0 10 Space velocity, \v.iiir./w 1. 58 3. 25 1. 3. 19 3. 22 l. 54 3.07 3.00 1. 46 Catalyst/oil ratio, w./w 7. 59 3. 7. 51 3. 76 3. 72 7. 3. 91 4. 00 8. 23 Product yields:

Ttztaii conversion (vol. percent of 46. 25 35. 68 48. 94 33. 70 35. 99 47. 04 42. 02 52. 59. 33

ee Gasolincs, 05-430" F. (vol. percent of c 29. 80 25. 32 31. 37 27. 36 28. 31 32. 55 33. 27 40. 44 40. 36 Cycle oil, 430IF. (v01. percent of iced) 53. 74 64. 32 51. 06 66. 30 64. 01 52. 96 57. 98 47. 15 40. 67 Coke (wt. percent of feed). 5. 85 3.68 6. 22 2. 58 2.22 4. 46 2. 74 2. 86 4. 76

ammonia. The hydrofining cond1t1ons and specifications TABLE HLB of the hydrofined 011s are tabulated in Table II. [zeolitic aluminosilicate catalyst] Run No 10 11 12 13 TABLE II Nitrogen in feed, p.p.m 4, 650 3, 200 1, 700 470 Run No A B C D 55 Operating conditions:

Reactor temperature, F 950 910 910 910 Operating cond1t1ons: Reactor pressure, p.s.i.g 10 i0 10 10 Temperature. F 6 4 5 7 7 Space velocity, w./hr./w 3. 21 3.15 3.07 3.00 Pressure, p.s.i.g. 1,000 Catalyst/oil ratio, w.lw 3. 74 3.80 3. 3. 99 Recycle gas T3120, S.0.1'./bbl 3, 450 Product yields; H2 Pressure, l 710 Total conversion (vol. percent of H: consumption, s.c.f./bbl 7 0 feed 5. 03 43. 60 57. 44 68. 87 Inspections: Gasolines, C 430 F. (vol. percent Gravity, API 20.3 20- 9 22- 5 2 -9 of fee 33. 74 33. 50 43. 90 53. 62 l qp F 5 150 146 149 Cycle oil, 430 F.+ (vol. percent of Basic n trogen, p.p.m. 1,330 1. 550 5 feed 54. 97 56. 40 42. 56 31.13 T015111 nltmgen, D-D- 200 500 700 470 Coke (wt. percent of iced) 5. 00 4. 10 4. 63 4. 58 Sulfur, wt. percent" 0.14 0. 08 0.03 0. 024 Metals, p.p.rn.:

V 0. 01 0. 01 0. 01 65 NL. 0.15 0. 01 0. 01 F0 Nil il Nil Substantially all the sulfur (i.e. approximately 90%) was removed by the mild hydrofining conditions of Run A. However, only about 30% of the nitrogen (3200 ppm. nitrogen remaining in the hydrofined feed) was removed in Run A. More severe hydrofining conditions (Run B) were required to reduce the nitrogen content of the gas oil to about 2500 p.p.m.

Using a feed having the same nitrogen content, and at the same reaction conditions, catalytic cracking with the catalyst comprising a zeolitic aluminosilicate produced higher yields of lower boiling products, including gasoline, than catalytic cracking with the amorphous silica-alumina catalyst. However, the increased conversion of the feed to lower boiling products using the zeolitic aluminosilicate catalyst was accompanied by increased coke production 9 (compare, for example, Run 12, Table III-B with Run 7, Table IIIA). At a constant level of coke production, conversion to lower boiling products, including gasoline,

V was substantially the same for the two catalysts with high nitrogen feeds (compare, for example, Run 10, Table III-B with Run 1, Table III-A), whereas conversion was substantially higher with the zeolitic aluminosilicate catalyst when using low nitrogen (hydrofined) feeds (compare, for example, Run 13, Table III-B with Run 9, Table III-A) Since most commercial catalytic cracking operations are limited by their coke-burning capacity, it is preferable to compare the effectiveness of different catalysts for catalytic conversions at a constant coke yield. FIGURE 1 shows, for comparative purposes, the catalytic cracking conversion (total conversion) as a function of nitrogen content at a constant coke yield of 5 weight percent using a zeolitic aluminosilicate containing catalyst and the amorphous catalyst. Using hydrocarbon oils containing high nitrogen levels, that is, greater than about 2500 ppm. nitrogen, the conversion to lower boiling products for the two catalysts was approximately the same. However, as the nitrogen content of the hydrofined feed was decreased below about 2500 p.p.m., conversion at constant coke deposition increased significantly for the zeolitic aluminosilicate catalyst in relation to the amorphous catalyst. The smaller the nitrogen content of the feed (the more effective the hydrofining operation to remove nitrogen compounds), the greater the dilference between the catalytic cracking conversion for the two catalysts at a constant coke deposition. Thus, it is preferred, prior to catalytic cracking, to reduce the nitrogen content of a feed having a nitrogen content greater than 2500 p.p.m. to below 2500 p.p.m. and preferably below 2000 p.p.m. and more preferably below at least 1000 ppm.

As the nitrogen content of the feed is lowered, the catalytic cracking conversion of the process using the zeolitic aluminosilicate increases more rapidly than the catalytic cracking conversion of the process using the amorphous catalyst. Thus, referring to FIGURE 2, it is evident that hydrofining a feed containing, for example, 2500 p.p.m. nitrogen, to reduce the nitrogen content to, for example, 2000 p.p.m., then catalytically cracking the hydrofined feed with a zeolitic aluminosilicate catalyst, results in an incremental advantage over a similar hydrofining-catalytic cracking combination using an amorphous catalyst in the catalytic cracking step. Thus, the incremental advantage of catalytic cracking of a feed containing 2500 p.p.m. with a zeolitic aluminosilicate catalyst compared to catalytic cracking with an amorphous catalyst is only about 1.8 percent (the difference, in conversion, between point b and point a of FIGURE 2). However, at 2000 p.p.m. nitrogen the incremental advantage of catalytic cracking with the zeolitic aluminosilicate catalyst over catalytic cracking with an amorphous catalyst is about 3.8 percent .(the difference, in conversion, between point d and point c of FIGURE 2). Similarly, hydrofining a feed containing, e.g., 2000 p.p.m. nitrogen to reduce the nitrogen content to, e.g., 1500 p.p.m., then catalytic cracking, results in an unexpected increase in catalytic conversion using the zeolitic aluminosilicate catalyst compared to the amorphous catalyst. At 2000 p.p.m., the advantage of a zeolitic aluminosilicate catalyst over an amorphous catalyst is 3.8 percent compared to 5.6 percent at 1500 p.p.m. Table IV shows the significant difference in the catalytic cracking conversion of feeds of different nitrogen content between processes using the zeolitic aluminosilicate catalyst and processes using the amorphous catalyst. The data in Table IV were obtained from FIGURE 2; the points on the curves in FIGURE 2 at which the difference was determined are depicted by letters connected with a dashed line.

Thus, it is apparent that an unexpected advantage is obtained by hydrofining a feed containing substantial amounts of organic nitrogen, e.g., 500 to 2500 p.p.m. to lower the nitrogen content substantially, e.g., by several hundred p.p.m., and then catalytically cracking the hydrofined feed with a crystalline zeolitic aluminosilicate catalyst.

What is claimed is: '1. A process for catalytically cracking a high boiling hydrocarbon charge stock having a nitrogen content above 1000 p.p.m., comprising:

contacting the charge stock and from 1000 to 4000 s.c.f./bbl. of hydrogen with a hydrodenitrification catalyst at a pressure in the range from 800 to 3000 p.s.i.g., at a temperature in the range from 600 to 800 F., and at a liquid hourly space velocity in the range from 0.3 to 5.0 volume of feed per volume of catalyst in a hydrofining zone to reduce the nitrogen content of the charge stock to below 500 p.p.m.;

deammoniating the effluent from the hydrofining zone;

and then contacting the substantially ammonia-free reduced nitrogen content charge stock under catalytic cracking conditions, with a cracking catalyst comprising a crystalline zeolitic aluminosilicate having pore diameters of at least 6 angstroms. 2. A process for catalytically cracking a high boiling hydrocarbon charge stock having a nitrogen content above 25 00 p.p.m., comprising:

contacting the charge stock and from- 1000 to 4000 s.c.f./-bbl. of hydrogen with a hydrodenitrification catalyst at a pressure in the range from 800 to 3000 Y p.s.i.g., at a temperature in the range from 600 to 800 F., and at a liquid hourly space velocity in the range from 0.3 to 5.0 volume of feed per volume of catalyst in a hydrofining zone to reduce the nitrogen content of the charge stock to below 500 p.p.m.;

deammoniating the effluent from the hydrofining zone;

and

then contacting the substantially ammonia-free reduced nitrogen content charge stock, under catalytic cracking conditions, with a cracking catalyst comprising a crystalline zeolitic aluminosilicate having pore diameters of at least 6 angstroms.

3. A process as in claim 2 wherein the crystal ine zeolitic aluminosilicate is of the Y crystal type.

References Cited UNITED STATES PATENTS 2,962,435 11/1960 Fleck et al. 208-119 3,272,734- 9/ 1966 McLaren 208-89 DELBERT E. GANTZ, Primary Examiner A. RIMENS, Assistant Examiner US. 01. 'X.R. 20s 1zo 

